Process for the polymerization of monomers

ABSTRACT

THERE IS PROVIDED A PROCESS FOR THE COPOLYMERICATION OF ETHYLENE IN A HORIZONTAL CYLINDRICAL EVAPORATIVELY COOLED REACTOR PROVIDED WITH A ROTATABLE RIBBON BLENDER AGITATOR HAVING A SET OF OUTER RIBBON BLADES PITCHED IN A DIRECTION OPPOSITE TO A SET OF INNER RIBBON BLADES.

Dec. 19, 1972 E. P. GOFFINET, JR 3,795,119 I PROCESS FOR THE PQLYMERIZATION 0F MONOMERS Filed 001:. e, 1971 CV 30 1 T I T 32 CV INVENTOR EDWARD P. GOFFBNET, JR.

BYCEVMZW ATTORNEY United States Patent Office Patented Dec. 19, 1972 3,706,719 PRGCESS FOR THE POLYMERIZATION F MGNOMERS Edward P. Goftinet, Jr., Wilmington, DeL, assignor to E. 1. du Pont de Nemours and Company, Wilmington, Del.

Filed Oct. 6, 1971, Ser. No. 187,080 int. Cl. C08f /40 US. Cl. 260-8058 4 Claims ABSTRACT OF THE DISCLOSURE There is provided a process for the copolymerization of ethylene in a horizontal cylindrical evaporatively cooled reactor provided with a rotatable ribbon blender agitator having a set of outer ribbon blades pitched in a direction opposite to a set of inner ribbon blades.

This invention relates to continuous polymerization of ethylene, and a C -C alpha-olefin in an evaporatively cooled reactor.

The polymerization of monomers, such as ethylene, with one or more alpha-olefins, such as propylene, in solution using a coordination catalyst system is well known. The polymerization can be conducted continuously to obtain an ethylene copolymer, dissolved in a suitable solvent, which is isolated by conventional procedures. The most efficient of such processes are those which are based on increasing the concentration of polymer in the reactor in order to reduce the amount of solvent which has to be removed during isolation of the polymer product. How ever, an increase in polymer concentration leads to a relatively large increase in the solution viscosity. This in turn aggravates two problems: (1) the rate of mass transfer of ethylene from the gas to the liquid phase (i.e., ethylene mass transfer coefficient) is reduced, and (2) the rate of heat transfer to remove the large exothermic heat of polymerization is reduced.

The heat transfer problem can be solved economically by employing evaporative cooling wherein reaction heat is removed by evaporating ethylene and propylene monomers from the solution during the reaction and cooling and recycling them to the reaction mixtures. An evaporativelycooled reactor has the disadvantage, however, that ethylene concentration in solution in the reactor is usually less than its equilibrium value. In order to produce a copolymer containing a particular proportion of ethylene, it is usually necessary to recycle a larger amount of ethylene in the reactor off-gas than would be the case if concentration of ethylene in solution in the reactor achieved its equilibrium value. Economically, this increase in recycle volume means greater expense than would otherwise be the case. There has been a need for a process for copolymerizing ethylene in an evaporatively-cooled reactor whereby improved mass transfer of ethylene is provided and ethylene concentration in solution in the reactor is maintained at a value close to its equilibrium value even when the solution has a viscosity greater than about 500 centipoise (cp.).

There is provided by this invention an improved process for the solution polymerization of ethylene and at least one C -C alpha-olefin in the presence of a coordination catalyst in an agitated evaporatively cooled reactor containing both a liquid and a vapor phase, the liquid phase having a solution viscosity greater than 500 cp. at the reaction conditions wherein at least a part of the off-gas from the reactor is compressed and condensed to form a liquid condensate, the condensate is partially vaporized to form a mixture of liquid and gas, and this mixture is recycled to the reactor. The improvement involves agitating the gas and liquid phases in a generally cylindrical reaction vessel having a horizontally aligned axis and a helical blade type agitator with at least two coaxial helical blades having oppositely directed pitches (reverse pitched blades) mounted coaxially on and at diflerent radial distances from a rotatable horizontal shaft extending along the axis of the cylinder, the reaction vessel (reactor) being about 35- filled with liquid reaction mixture and the agitator speed being sufiicient to provide an ethylene mass transfer coefiicient equivalent to at least about 0.3 min. at a solution viscosity of less than about 5000 centipoise for a reactor of at least 5000 gallon (18,950 liters) capacity.

Referring to FIG. 1, which depicts a preferred embodiment of the reaction vessel utilized in the invention, horizontal reactor 14 is shown in cross-section, with cylindrical heads 19, vapor outlet port 3, product removal port 12, and inlet ports 1, 2, 4, 10, 11, and 13 for introduction of reactants and other reaction mixture components. The inlet ports can be fitted with dip tubes or sparge tubes to aid in distribution of feed reactants. The reactor can be constructed of materials well known in the art, such as carbon steel and stainless steel.

Reactor 14 is provided with a spiral agitator comprising a set of inner helical blades 15 pitched in a direction opposite to a set of outer helical blades 16 which move close to the reactor wall. Both sets of blades are mounted coaxially on a rotatable horizontal shaft 17. Members 18 provide support for the spirals. The agitator blades, shaft and supports can be fabricated of materials commonly used in the art. Typically, blades 15 and 16 are made from carbon or stainless steel, cast or machined to the proper shape; and welded 0r bolted to the shaft 17.

The polymers produced by the process of this invention are hydrocarbon soluble copolymers of ethylene (E) containing about 20-60% by weight of C -C alpha-olefin units, such as propylene (P) units. Preferred are copolymers of ethylene containing about 25-50% propylene and l-20% nonconjugated diene units by weight and preferably about 25-50% propylene and 2-10% diene by weight. By the term hydrocarbon soluble copolymers is meant that copolymer formed in the reactor is soluble at about -20 C. to about C. in the polymerization solvent used in the reactor. The non-conjugated hydrocarbon diene utilized contains about 5-22 carbon atoms. Representative dienes are 1,4-hexadiene, dicyclopentadiene, S-methylene 2 norbornene, S-ethylidene-Z-norbornone, and 1,5-cyclooctadiene. The general preparation of the copolymers produced by this invention are well known.

The process of this invention is carried out in a hydrocarbon solvent, preferably in a saturated linear hydrocarbon solvent containing about 5-8 carbon atoms such as pentane, hexane, and heptane. Compatible mixtures of solvents can also be employed. A coordination catalyst system of the type now well known is utilized. It should be noted that for each different catalyst combination a different P/E mole ratio in the solution in the reactor may be required to yield a given propylene content in the polymer. For each catalyst system there will exist a certain lower limit of the P/E mole ratio for the liquid monomers in the polymerization solution at which will be formed copolymer sufficiently high in ethylene content to be insoluble in the solvent. Preferred catalyst systems consist essentially of trisacetylacetonate (VAA)/ diisobutyl aluminum monochloride (DIBAC); vanadium tetrachloride (VCl )/DIBAC; vanadium oxytrichloride (VO'Cl )/dietl1yl aluminum ethoxide/DIBAC/benzotrichloride and VOCl /DIBAC. In using such catalysts, suitable aluminum to vanadium (Al/V) mole ratios and amounts of catalyst employed per liter of total solution will depend on the specific compounds and conditions employed. These are well known. Usually, Al to V mole ratios will fall within the range from about 2:1 to about 20:1. The amount of catalyst expressed as the amount of vanadium content usually ranges from about 0.00001 to 0.002 mole per liter of polymerization reaction mixture.

Coordination catalyst compositions containin compounds which enhance their activity, such as benzotrichloride, hexachloropropene, and the like can be employed. In using these adjuvants the Al/V ratio may be greatly increased and the vanadium concentration greatly reduced.

The temperature at which the polymerization reactor is operated will have some effect on the maximum amount of ethylene which can be tolerated in the polymer without the polymers becoming insoluble in the polymerization solvent. While the polymerization reactor can be operated satisfactorily at temperatures rangingfrom about -20 C. to about +80 C., it is preferred that the reaction be conducted within the range of about 20 C. to about 60 C. Generally, the pressure is at least about 30 p.s.i., and preferably less than about 200 p.s.i. The pressure is dependent on the temperature and monomer concentrations in the reaction medium.

Ethylene can be fed to the reactor both as a vapor and as a liquid along with propylene in such amounts that the ratio of liquid propylene to liquid ethylene being fed to the reactor is high enough to produce a soluble polymer at the inlet port. There is no upper limit for P/E mole ratio in the liquid feed insofar as the elimination of insoluble polymer is concerned, but the P/E in the liquid feed need not be higher than the P/E mole ratio in the reactor liquid phase. A practical determination of the minimum P/E mole ratio in the liquid fed to the reactor under a given set of conditions can be made by setting up the continuous process with the desired amount of recycle of monomers stripped from the product, recycle of reactor off-gas, and make-up monomers, and running it with decreasing P/E mole ratios in the liquid feed to determine at which point substantial reactor inlet fouling occurs. It is preferred that the liquid being fed to the reactor have a P/E mole ratio of at least 3:1.

The process of the invention can be better understood by referring to FIG. 2 which is described below with respect to copolymerization of ethylene, propylene and 1,4-hexadiene. Vanadium tetrachloride (VCI and diisobutyl aluminum monochloride [(i-C H AlCl] are fed continuously as separate streams 21 and 22 to reactor A containing hexane solvent where the catalyst is formed in situ. 1,4-hexadiene is also continuously fed as stream 33 to reactor A. Polymer product, consisting essentially of copolymer in solution, unreacted monomers and catalyst, is continuously removed via stream 32 through control valve CV which throttles the exit flow rate and maintains liquid volume in the reactor at approximately one-half the total volume. A portion of the more volatile components of the reaction mixture is withdrawn continuously as vapor stream 23 through a flow control valve CV; which is regulated by temperature sensor T. The vapor is passed through entrainment separator B, and the liquid portion returned to the reactor as stream 25. Gas in stream 24, is fed to compressor C, where it is compressed and then combined with fresh ethylene feed from stream 26, and fresh propylene feed from stream 27, before entering condenser D for partial condensation. The combined gas-liquid mixture from the condenser passes as stream 28 to control valve CV where it is flashed into gas-liquid separator (E). Gas and liquid, streams 30 and 31, respectively, are returned to the reactor through nozzles placed in the liquid portion of the reactor. Inert gas is eliminated as stream 34.

Y The latent heat of vaporization of the recycle liquid and the sensible heat capacity of the vapor-liquid mixt-ure provide a heat sink for the removal of the exothermic heat of polymerization. This method of cooling is referred to as evaporative cooling. Operating temperatures,

pressures, stream compositions, and flow rates for the various streams shown in FIG. 2 are listed in Table 2. These operating conditions reflect the material balance, energy balance, reaction kinetics, and mass transfer characteristics of a vertical reactor for the production of 10,000 pounds of polymer per hour. The numerical values in Table 2 are calculated values since they can be calculated with greater precision than they can be measured. However, they are representative of conditions established in actual practice.

The reactor total volume is nominally 16,500 gallons (62,500 liters) based on a cylindrical vessel 11 feet (3.4 meters) diameter and 22 feet (6.7 meters) in height. The ungassed liquid volume in the reactor is 8250 gallons (31,200 liters); the liquid having a specific gravity of 0.66. The reactor is agitated by a conventional turbine type agitator with a vertical shaft and a speed of rpm. and is also fitted with four equally spaced, radially oriented, vertical bafiies attached to the wall of the reactor. Suitable ports for introduction of gas and liquid feeds and removal of liquid product are provided at the bottom of the vessel, and suitable ports are located at the top of the reactor for the removal of vapor.

Under the operating conditions listed in Table 2, the reaction mixture solution viscosity is 1250 cp. At this viscosity, the experimentally determined ethylene volumetric mass transfer coetficient, K a, is 0.28 min.- in accordance with the following equation:

Km: 1Ethylene fllux .TP mFw/ h) where:

This equation can be used: (1) to correlate actual data on reactor operating conditions such as that shown in Table 2 to obtain a value of (K a) useful for design purposes; and (2) to design a reactor for new conditions in which (P/E) is calculated from known values of all other factors including (K a).

The following examples illustrate these two possibilities:

(1) Calculation of (Km) from reactor operating conditions In the example shown in Table 2, the polymer production rate is 10,000 lb./hr. (4540 kg./hr.) and the ethylene content of the polymer is 64 weight percent. The average molecular weight of the polymer-free reactor solution (stream 32 after separation of polymer) is 91,640- 10,000 1,322-290.5 The ethylene flux is:

= 79. 1 lb. /hr. mole polymer-free solution Ethylene fiux= The relative volatility of ethylene to propylene at 30 C. is 4. From Table 2, stream No. 32,

)=0.00665 min.-

6 and the value of In order to obtain an improved rate of mass transfer 1432 of ethylene from the vapor to the liquid phase in a re- (P/E)1,= =7.96 actor containing ethylene and propylene, it is necessary to observe certain design criteria for the reactor and For Stream 23, the Value of agitator. The length of the reactor should be two to four 18452 5 times its diameter. The choice of design parameters for )v= the inner and outer spiral agitator blades, shown in Table 1, are not independent of each other 'but are selected f fi i "g (KI-a) can be calculated so that the volumetric pumping capacity in the axial diquallon as 0 rection of the inner and outer sets of blades are aproximatel e ual and o osite. 0.00665=K1.a (0.139) p V q PP (4)(0.840) 7.96 TABLE 1 K 1441:0285: 10% [Preferred specifications for agitator] (2) Calculations of fOI 116W Outer set of Inner set of reactor conditions 18 blades blades Assume that (K a) i known for a new production Number Ofblades Aboutito Abmltito about 12. about 12. rate from previous correlations of the reactor Perform Blade width About 0.025 to About 0.04 to ance. The polymer production rate and its composition Dim er f it t H t m d about 9 .10- about 0.20. are chosen and from previously established correlations 20 fj 0 as a 0 er a 9 33? .2.? 533333;? Of the kinetics Of POlYIIlfiil'lZfltiOIl (P/E)L 1S fi d' The Pitch ofbladesa 1 4 2L to k. average molecular weight of the polymer-free solution 2 is then calculated from the reactor solution composition. p ed; widt h isleirtpressetll a; a frtctcilog of the vessel diameter.

e iame er c'r esc th 1' t t fth 't t i Thls leaves (P E )V to be obtamed from Equatlon measured to the uter etigi bi th ag itatgr bit-(1 2) z l f actig gfir s sel This is the actual procedure followed in calculating diarIneter. I h t d d B itch o t e agita or bla e is measure as the length L o! a vessel (P E V Whlch 18 shown m Table 2 as stream The required to cause one complete revolution of the agitator. The negative magnitude of the flow rates of each of the components sign on the pitch indicates that the pitch of the inner spiral is opposite in stream 23 is determined from the heat evolved in that the Spiral- TABLE 2 [Operating conditions for production of 10,0001b./hr. (4,540 kgJhr.) EPDM terpolymer in a vertical turbine-agltated reactor-Example 1] Stream Phase Liq. Liq. Gas Gas Llq. Gas Liq. Gas Llq. Gas Liq. Llq. Liq.

Flow rate, lb.-mo1eslhr.:

Nitrogen 91. 9 91. 9 87. 6 4. 3 91. 7 2 Ethy1ene 2 196. 6 2 195.7 0.9 247. 6 l, 624. 6 813. 7 2 323. 4 119. 2 18.0 Propylene-.- l 5. 2 1 1. 3 3. 9 .3 1, 213. 7 l 5. 5 465. 5 Hexane 4. 4 7. 4 37. 4 39.). 0 0. 8 396. 6 0 2 Temp. C 1

Pressure (p.s.i.g.)- 150 150 25 25 25 450 450 450 450 e5 e5 25 150 forming the polymer and the heat capacities and heat The invention permits preparation of copolymers,

of evaporation of unit amounts of each of the components especially ethylene/propylene copolymers, whose utility in stream 23. The fraction of the gas from stream 23 as valuable elastomers is well known, by an improved which is condensed in the partial condenser (D), deprocess which affords increased reactor capacity with a pends on the temperature, pressure, and composition of minimum of reactor fouling. Polymer of excellent quality the compressed vapor. For the conditions shown in can be produced with good catalyst efliciency.

Table 2, the fraction condensed in (D) is 48 mole per- The operation of the process of this invention is illuscent. The required recycle coolant flow rate is 18 lb. of trated by the following examples in which all parts, proreactor off-gas per pound of polymer produced. portions, and percentages are by Weight unless otherwise In operating the vertical reactor, as described above, indicated.

the reactor fouls frequently, averaging once per day, EXAMPLEl because the condensate returned to the reactor is overly vertical reactor with a turbhwtype agitator rich in its of ethylene. AS shown m Table thls A conventional vertical, cylindrical, turbine-agitated condensate m Stream 31 comams D1916 Percent t i' reactor is used in accordance with the process shown in we 3 mole Percent Pmpylene yleldmg a 2 FIG. 2 for the copolymerization of ethylene, propylene the l1qu1d feed of 3.9. At a reaction temperature of 30 and 1,4 hexadiene in solution to produce an 11 weight C. such a P/E value causes reactor fouling by producpercent solutlon of ethylene/propylene/1,4-hexad1ene mg insoluble products. To cope with such reactor fouling, (EPHD) terpolymer in the reactor emuent. The p y operation of the reactor must be stopped and the reactor met contains 32 percent propylene, 64 percent ethylene, Z5325? 2x5352392 g i i g gg g ggsg gg i ai f and 4 percent 1,4-hcxadiene; the molecular weight of the creasing the ethylene content and increasing the propylene g, j zssmeasured by Mooney vlscoslty (ML-4 at content of the condensate from condenser D. Fouling EXAMPLE 2 is more conveniently prevented in accordance with the instant invention, by using a horizontal reactor as de- Horlzontal reactor Wlth a S-PltCh spiral-type scribed below, whereby a greater ethylene mass transfer agltator coefiicient is realized with the result that the reactor A horizontal cylindrical vessel with a double spiral vapors in stream 23, are relatively richer in propylene. reverse-pitch agitator as shown in FIG. 1 is used for 7 comparative purposes to polymerize the same monomers in the same proportions as described in Example 1, using the process of FIG. 2. The reactor total volume is 16,500 gallons (62,500 liters) based on a cylindrical vessel The precision of the coefiicients in Table 3 is i40%. For the plant conditions specified in Table 4, the reactor solution viscosity in stream 12 is 1250 op. The ethylene mass transfer coeflicient for the horizontal renominally 11 feet (3.3 meters) in diameter and 22 feet 5 actor is 1 minr which is almost four times greater than (6.6 meters) long. The ungassed liquid volume is 8,250 the value (0.28 minr measured for the vertical turbinegallons (31,300 liters). The agitator consists of an outer agitated reactor. set of ei ht ribbon blades and an inner set of six ribbon blades nfounted coaxially on a horizontal shaft running Advantage of the honzpntal reactor compared to the length of the reactor. The outer set of eight blades are the Vemcal reactor equally spaced on a 360 circle with an outer diameter of The propylene content of the ofi-gas in stream 23 from ten feet (3 meters). Each of the outer blades is six inches the horizontal reactor is 50.9 mole percent compared to (0.15 meter) wide and makes one-eighth turn over the only 40.2 mole percent from the vertical reactor. The length of the vessel. The inner set of six blades are equal increased content of propylene in the reactor off-gas ly spaced on a 360 circle with an outer diameter of 5.5 means that 65 mole percent of the compressed vapor feet (1.8 meters). Each of the inner blades is nine inches from (C) can be condensed in (D) for the same com- (22.8 cm.) wide and makes one full turn in the length pressor and condenser operating conditions as used for of the vessel but in the opposite direction from the outer the vertical reactor in Example 1. As a result, the coolant blades when viewed from either end of the vessel. Agitator flow rate, stream 23, is reduced to 13 pounds of vapor speed is 35 r.p.m. per pound of polymer. The propylene content of the Suitable ports are provided in vessel 14, FIG. 1, for: condensate produced from the condenser (D) in stream (a) inlets for gas and liquid feeds below the liquid sur- 28 (Table 3) is 57.4 mole percent for the horizontal reface (1, 2, 4, 10, 11, and 13); (b) seals for the agitator actor as compared to 49.8 mole percent for the vertical shaft; (c) an outlet for the vapor 3; and (d) an outlet for reactor (stream 28 in Table 2). Fouling of the horizontal the polymer solution 12. reactor is practically eliminated because of the decreased The operating temperatures, pressures, stream comethylene content of the condensate from the condenser positions and flow rates for this reactor system are pre- (D). In addition, the improved mass transfer coefficient sented in Table 3. Reaction conditions are shown for is advantageous in reducing compressor size, investment the production of 10,000 lbs./hr. (4540 kg./hr.) of terand operating costs.

TABLE 4 [Operating conditions for production of 10,000 lb./hr. EPDM terpolymer in a horizontal agitated reactor-Example 2] Stream 21 22 23 24 25 2s 27 2s 29 so 31 32 33 Phase Liq. Liq. G88 G85 Llq Gas Liq. Gas Liq. Gas Liq. Liq. Liq.

Flow rate, lb. moles/hm Nit VCl-i (l-CAI'IQ)2AIC1 Total, lb. molcs/hr 418. 7 442. 5 3, 181 3, 144 36. 6 247. 6 219. 7 1, 250 2, 362 2, 498 1, 114 1, 322 14. 1

Lbs/hr 36, 003 38, 201 130, 407 127, 474 2, 933 6, 044 9, 244 42, 114 101, 548 86, 374 57, 288 01, 640 1, 158 Temperature. 10 10 30 30 25 25 40 40 -14. 4 14. 4 30 25 Pressure (p.si.g.) 150 150 25 25 25 450 450 450 450 65 65 25 150 TABLE 3 Ethylene mass transfer coefficient in the above horizontal reactor with a double-spiral reverse pitch agitator Ethylene mass transfer Viscosity (cp.): coefilcient, K a (minr til What is claimed is:

1. In a continuous process for solution polymerization of ethylene and at least one C -C alpha-olefin in the presence of a coordination catalyst system in an evaporatively cooled reactor containing a vapor phase and a liquid phase with a solution viscosity greater than about 500 centipoises wherein vapor is removed from the reactor and partially condensed to form a liquid/vapor mixture which is recycled to the liquid phase of the reactor, the improvement of agitating the vapor and liquid phases in a horizontal reactor by revolving therein an agitator having a horizontal rotatable shaft and at least two reversepitched coaxial helical blades attached to said shaft at different radial distances therefrom, the reactor being about 35-75% filled with liquid reaction mixture and the agitator speed being suflicient to provide an ethylene mass transfer coefficient equivalent to at least about 0.3 min.- at a solution viscosity less than about 5000 centipoises for a reactor capacity of at least 5000 gallons.

2. The process of claim 1 in which the alpha-olefin is propylene.

10 3. The process of claim 2 in which ethylene, propylene 3,330,818 7/1967 Derby 260-949 and a non-conjugated diene are polymerized. 3,524,730 8/ 1970 Yokouchi 23-285 4. The process of claim 3 in which the diene is 1,4- 02 3/1971 Christensen 23-285 hexadlene' 5 JAMES A. SEIDLECK, Primary Examiner R. S. BENJAMIN, Assistant Examiner References Cited UNITED STATES PATENTS 2,963,470 12/1960 Lanning 26088.2 Cl 3,300,457 1/1967 Schmid 260-882 1 

